Integrated process for making propene oxide and an alkyl tert-butyl ether

ABSTRACT

An integrated process for making propene oxide and an alkyl tert-butyl ether comprises dehydrogenating a feed stream comprising iso-butane to provide a stream comprising iso-butene and hydrogen and separating this stream into a stream consisting essentially of hydrogen and a stream comprising iso-butene; reacting a part or all of the stream comprising iso-butene with an alkanol in the presence of a solid acid catalyst to provide an alkyl tert-butyl ether; reacting a part or all of the stream consisting essentially of hydrogen with oxygen, providing a stream comprising hydrogen peroxide; and reacting a part or all of the stream comprising hydrogen peroxide with propene in the presence of an epoxidation catalyst to provide propene oxide.

FIELD OF THE INVENTION

The present invention is directed to a an integrated process for makingpropene oxide and an alkyl tert-butyl ether, which makes efficient useof the feed materials and at the same time allows a variation in theratio of propene oxide to alkyl tert-butyl ether product over a broadrange.

BACKGROUND OF THE INVENTION

A coupled production of propene oxide and methyl tert-butyl ether (MTBE)is known from the prior art using the so called Halcon process, whereiso-butane is oxidized with air to tert-butyl hydroperoxide, which isthen reacted with propene in the presence of a molybdate catalyst togive propene oxide and tert-butanol. The tert-butanol is further reactedwith methanol to give MTBE. However, such a coupled production, whereboth products are obtained at an essentially fixed ratio, has adisadvantage when market demand for the two products does not match thisproduct ratio. Since the demand for propene oxide has grown more rapidlyin recent years than the demand for MTBE, the HPPO process for makingpropene oxide from propene and hydrogen peroxide has been developed as astand-alone process for making propene oxide. However, separateproduction lines for MTBE and propene oxide by the HPPO process requiremore raw material than the coupled production of propene oxide and MTBEby the Halcon process.

SUMMARY OF THE INVENTION

The inventors of the present invention have now found a way ofintegrating the hydrogen peroxide based manufacture of propene oxidewith the production of MTBE which reduces raw material consumptionsimilar to the Halcon process but at the same time remains flexible withregard to the production ratio of both products.

Subject of the invention is therefore an integrated process for makingpropene oxide and an alkyl tert-butyl ether comprising the steps:

-   a) a step of dehydrogenating a feed stream S1 comprising iso-butane,    providing a stream S2 comprising iso-butene and hydrogen;-   b) a separation step separating stream S2 into a stream S3    consisting essentially of hydrogen and a stream S4 comprising    iso-butene;-   c) a step of reacting a part or all of stream S4 with an alkanol in    the presence of a solid acid catalyst, providing a stream S5    comprising an alkyl tert-butyl ether;-   d) a separation step separating the alkyl tert-butyl ether from    stream S5;-   e) a step of reacting a part or all of stream S3 with oxygen,    providing a stream S6 comprising hydrogen peroxide;-   f) a step of reacting a part or all of stream S6 with propene in the    presence of an epoxidation catalyst, providing a stream S7    comprising propene oxide; and-   g) a separation step separating propene oxide from stream S7.

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 shows a flow chart for an embodiment of the integrated processwhere the feed stream to step a) additionally comprises n-butane and astream consisting essentially of 1-butene is obtained as an additionalproduct.

FIG. 2 shows a flow chart for an embodiment of the integrated processadditionally comprising a step of isomerizing n-butane to iso-butane toprovide the feed stream to step a).

DETAILED DESCRIPTION OF THE INVENTION

In step a) of the integrated process of the invention, a feed stream S1comprising iso-butane is dehydrogenated to provide a stream S2comprising iso-butene and hydrogen.

Feed stream S1 may be a butanes fraction or an iso-butane fractionseparated by distillation from a liquefied petroleum gas or from the lowboiling products of a catalytic hydrocracking process. Preferably, feedstream S1 is provided by an additional step of isomerizing n-butane toiso-butane. Suitable methods for isomerizing n-butane to iso-butane areknown from the prior art and are available for license, such as theButamer™ process of UOP. In the Butamer™ process, isomerization ispreferably carried out at a temperature of 180 to 220° C. and a pressureof 15 to 30 bar by contacting a gas stream containing butane andhydrogen at a molar ratio of from 0.5 to 2 with a platinum catalystsupported on alumina and activated with chloride. The reaction iscarried out by passing the gas stream over a catalyst fixed bed. Achlorinated organic compound is preferably added to the gas stream tomaintain chloride activation of the catalyst.

The product stream resulting from the isomerization reaction ispreferably separated into an n-butane and an iso-butane fraction, a partor the entire n-butane fraction is recycled to the isomerizationreaction and the iso-butane fraction or a mixture of the iso-butanefraction and the non-recycled n-butane fraction is provided as streamS1. Separation of the product stream into an n-butane and an iso-butanefraction can be achieved by distillation. Adjusting the separationefficiency of a distillative separation or varying the ratio of n-butanefraction recycled to isomerization to n-butane fraction passed withstream S1 to step a) allows to adjust the ratio of iso-butane ton-butane in feed stream S1 as needed.

Suitable methods for dehydrogenating a stream comprising iso-butane toprovide a stream comprising iso-butene and hydrogen are known from theprior art and are available for license, such as the Catofin™ process ofCB&I Lummus or the Oleflex™ process of UOP. Dehydrogenation ofiso-butane can be carried out at temperatures of from 500 to 700° C. andpressures of from 0.1 to 2.5 bar in the presence of a dehydrogenationcatalyst. Suitable dehydrogenation catalysts are chromium oxidesupported on alumina and activated with an alkali metal, used in theCatofin™ process, and platinum, promoted with tin and alkali metals,supported on alumina, used in the Oleflex™ process. The Catofin™ processis carried out with several adiabatic fixed bed reactors operated inparallel, alternating reactor operation between endothermaldehydrogenation and exothermal catalyst regeneration with air. In theCatofin™ process, dehydrogenation is preferably carried out at 580 to660° C. and 0.1 to 0.6 bar. The Oleflex™ process is operated with aseries of adiabatic moving bed reactors with gas heating betweenreactors and external catalyst regeneration. In the Oleflex™ process,dehydrogenation is preferably carried out at 550 to 650° C. and 1.0 to2.5 bar.

In step b) of the integrated process of the invention, stream S2obtained in step a) is separated into a stream S3 consisting essentiallyof hydrogen and a stream S4 comprising iso-butene. Stream S2 ispreferably cooled and compressed to condense hydrocarbons, providingstream S3 as a gas phase rich in hydrogen. The gas phase can be furtherpurified, preferably by a pressure swing adsorption. Methods forpurifying hydrogen by pressure swing adsorption are known from the priorart. The condensed hydrocarbons may be passed as stream S4 to step c).In a preferred embodiment, the condensed hydrocarbons are purified bydistillation, removing hydrocarbons having less than 4 carbon atoms asan overhead product and hydrocarbons having more than 4 carbon atoms asa bottoms product, to provide a stream S4 consisting essentially ofhydrocarbons having 4 carbon atoms.

In step c) of the integrated process of the invention, a part or all ofstream S4 obtained in step b) is reacted with an alkanol in the presenceof a solid acid catalyst to provide a stream S5 comprising an alkyltert-butyl ether. The alkanol is preferably methanol to provide methyltert-butyl ether (MTBE) or ethanol to provide ethyl tert-butyl ether(ETBE) and is most preferably methanol.

The solid acid catalyst is preferably an acidic ion exchange resin, morepreferably a resin having sulphonic acid groups. Suitable ion exchangeresins are sulphonated phenol/aldehyde condensates and cooligomers ofaromatic vinyl compounds. Examples of aromatic vinyl compounds forpreparing the cooligomers are: styrene, vinyltoluene, vinylnaphthalene,vinylethylbenzene, methylstyrene, vinylchlorobenzene, vinylxylene anddivinylbenzene. Preferably, the acidic ion exchange resin is asulphonated copolymer of styrene and divinylbenzene. The resin may be agel resin or a macroporous resin. Suitable ion exchange resins arecommercially available under the trade names Duolite® C20 and C26,Amberlyst® 15 and 35, Amberlite® IR 120 and 200,Dowex® 50 and Lewatit®SPC 118 SPC 108, K2611, K2621 and OC 1501, Amberlyst® 15, Amberlyst® 35and Lewatit® K2621 being preferred. The pore volume of the ion exchangeresins is preferably from 0.3 to 0.9 ml/g, especially from 0.5 to 0.9ml/g. The particle size of the resins is preferably from 0.3 mm to 1.5mm, especially from 0.5 mm to 1.0 mm.

The reaction is preferably carried out in a fixed bed reactor. When thealkanol is methanol, the reaction is most preferably carried out in twostages, using a liquid-filled fixed bed reactor in the first stage and areactive distillation column in the second stage.

In the first stage, stream S4 is preferably reacted with methanol at atemperature of 20 to 110° C., more preferably 25 to 70° C., and apressure of 0.5 to 5 MPa, more preferably 0.7 to 2 MPa, and temperatureand pressure are selected to maintain a liquid reaction mixture. Themolar ratio of methanol to iso-butene in stream S4 is preferably from10:1 to 1:1, more preferably from 5:1 to 1.1:1 and most preferably from1.8:1 to 1.2:1. The first stage is most preferably carried out with twoor three fixed bed reactors in series, operating the first reactor at 35to 70° C. and the second and optionally third reactor at 25 to 50° C.The residence time in the first stage is chosen to provide a compositionclose to the chemical equilibrium between methanol, iso-butene and MTBE,preferably achieving an isobutene conversion of more than 94%, morepreferably more than 96%. Tube bundle reactors and adiabatic fixed bedreactors may be used in the first stage.

In the second stage, the reaction mixture formed in the first stage isfurther reacted in a reactive distillation column, comprising an acidicion exchange resin arranged in a reaction zone, operated at a pressureof from 0.5 to 1.5 MPa, preferably 0.75 to 1.0 MPa, and a temperature inthe reaction zone of from 50° C. to 90° C., preferably from 55° C. to70° C., at a reflux ratio between 0.5 and 1.5, preferably between 0.7and 0.9. The reflux ratio refers to the ratio of the distillate streamreturned to the column to the distillate stream removed. The reactivedistillation column preferably has separation zones above and below thereaction zone. The separation zone above the reaction zone preferablyhas from 5 to 20 theoretical plates, in particular from 10 to 15theoretical plates. The separation zone below the reaction zonepreferably has from 12 to 36 theoretical plates, in particular from 20to 30 theoretical plates. The reaction mixture formed in the first stageis preferably fed to the reactive distillation column below the reactionzone, preferably 3 to 13 and more preferably 4 to 10 theoretical platesbelow the reaction zone. In the reaction zone, the ion exchange resinmay be integrated in a structured packing, for example in KataMax®packings known from EP 0 428 265, KataPak® packings known from EP 0 396650 or MultiPak® packings. Alternatively, it may be polymerized ontoshaped bodies as described in U.S. Pat. No. 5,244,929. Preference isgiven to using KataMax® packings. The amount of ion exchange resin ispreferably selected to achieve an isobutene conversion of 75 to 99%,preferably of 85 to 98% and more preferably of 95 to 97%, based on theisobutene content in the feed to the reactive distillation column.

In step d) of the integrated process of the invention, the alkyltert-butyl ether is separated from stream S5 provided in step c) asreaction product. The alkyl tert-butyl ether is preferably separatedfrom stream S5 by distillation. When the alkanol is methanol and step c)is carried out in two stages as described above, step d) may beintegrated with step c) in that the reaction mixture obtained from thefirst stage and from the reaction zone of the reactive distillationcolumn is separated in the separation zone below the reaction zone toprovide MTBE as the bottoms product of the reactive distillation columnand in the separation zone above the reaction zone to provide a mixtureof methanol and C4 hydrocarbons as the overhead product of the reactivedistillation column.

In a preferred embodiment of step d), a stream S8 comprising iso-butaneis additionally separated and from this stream S8 a stream S9 comprisingmore than 80% by weight iso-butane is separated and recycled to step a).The separation of streams S8 and S9 may be carried out by methods knownfrom the prior art, such as distillation, extractive distillation orextraction. When the alkanol is methanol and step c) is carried out intwo stages as described above, stream S8 is obtained as the overheadproduct of the reactive distillation column. Methanol is preferablyseparated from this overhead product by extracting with an aqueoussolution in a liquid-liquid-extraction. The aqueous solution preferablyhas a pH of at least 8, most preferably 8 to 12. The pH may be adjustedby adding a base, preferably sodium hydroxide. The extraction may becarried out in mixers and settlers or in an extraction column and ispreferably carried out in an extraction column operated in countercurrent. The extraction column has preferably 2 to 25 and morepreferably 5 to 15 theoretical plates and is preferably operated at atemperature of 10 to 90° C. and a pressure sufficient to maintain C4hydrocarbons as a liquid phase. The mass ratio of aqueous solution tothe overhead product of the reactive distillation column is preferablyfrom 1:5 to 1:40. If a feed stream S1 consisting essentially ofiso-butane is used in step a), extraction of methanol from stream 8 asdescribed above will provide a stream S9 comprising more than 80% byweight iso-butane which can be recycled to step a), optionally withfurther purification, such as drying.

In step e) of the integrated process of the invention, a part or all ofstream S3 consisting essentially of hydrogen is reacted with oxygen,providing a stream S6 comprising hydrogen peroxide. Additional hydrogenmay be fed to step e) in order to decouple the production capacity instep e) from the production capacity of steps a) to d).

Stream S3 can be reacted with oxygen in a liquid reaction medium in thepresence of a noble metal catalyst in what is known as a hydrogenperoxide direct synthesis. The noble metal catalyst is preferably asupported catalyst, with alumina, silica, titanium dioxide, zirconiumdioxide, zeolites and acticated carbons being preferred supports. Thenoble metal catalyst may be a suspended catalyst or preferably a fixedbed catalyst. The noble metal catalyst preferably comprises palladium asnoble metal, optionally in combination with platinum, gold or silver, acombination of palladium with platinum at a weight ratio of Pd:Pt ofmore than 4 being most preferred. Oxygen can be used as pure oxygen, airor air enriched in oxygen. Direct synthesis is preferably carried outwith a gas composition that is not flammable. For this purpose, an inertgas such as nitrogen or carbon dioxide can be added. Direct synthesis ispreferably carried out with a gas mixture containing at most 6% byvolume hydrogen and most preferably from 3 to 5% by volume hydrogen. Thegas mixture preferably contains preferably from 10 to 50% by volumeoxygen and most preferably from 15 to 45% by volume oxygen. Stream S3and oxygen are preferably dispersed separately in the liquid reactionmedium and inert gas can be added either to stream S3 or to the oxygenfeed. The liquid reaction medium may be an aqueous, aqueous-organic ororganic reaction medium and preferably consists essentially of analcohol or a mixture of an alcohol and water, the alcohol mostpreferably being methanol. The liquid reaction medium preferablycomprises a halide, more preferably iodide or bromide and mostpreferably bromide in an amount of 10⁻⁶ to 10⁻² mol/l, preferably 10⁻⁵to 10⁻³ mol/l and most preferably 10⁻⁵ to 5·10⁻⁴ mol/l in order tosuppress decomposition of hydrogen peroxide on the noble metal catalyst.The liquid reaction medium preferably further comprises a strong acidhaving a pK_(a) of less than 3 in an amount of 0.0001 to 0.5 mol/l andpreferably 0.001 bis 0.1 mol/l in order to improve selectivity forhydrogen peroxide formation, with sulfuric acid, phosphoric acid, nitricacid and methane sulfonic acid being preferred. The hydrogen peroxidedirect synthesis is preferably carried out in a fixed bed reactoroperated as bubble column with stream S3, oxygen and optionally inertgas being dispersed below a catalyst fixed bed.

In a preferred embodiment, stream S3 is reacted with oxygen in ananthraquinone process, providing a 20 to 75% by weight aqueous solutionof hydrogen peroxide. The anthraquinone process uses a working solutioncomprising at least one 2-alkylanthraquinone,2-alkyltetrahydroanthraquinone or a mixture of both, referred to asquinones in the following, and at least one solvent for dissolving thequinone and the hydroquinone. The 2-alkylanthraquinone is preferably2-ethylanthraquinone (EAQ), 2-amylanthraquinone (AAQ) or2-(4-methylpentyl)-anthraquinone IHAQ and more preferably a mixture ofEAQ with AAQ and/or IHAQ where the molar fraction of quinones carryingan ethyl group is from 0.05 to 0.95. The working solution preferablyfurther comprises the corresponding 2-alkyltetrahydroanthraquinones andthe ratio of 2-alkyltetrahydroanthraquinones plus2-alkyltetrahydroanthrahydroquinones to 2-alkylanthraquinones plus2-alkylanthrahydroquinones is preferably maintained in the range of from1 to 20 by adjusting the conditions of the hydrogenating andregenerating steps used in the anthraquinone process. The workingsolution preferably comprises a mixture of alkylbenzenes having 9 or 10carbon atoms as solvent for anthraquinones and at least one polarsolvent selected from diisobutylcarbinol (DiBC), methylcyclohexylacetate(MCA), trioctylphosphate (TOP), tetrabutylurea (TBU) andN-octylcaprolactam as solvent for anthrahydroquinones, DiBC, MCA and TOPbeing preferred and TOP being most preferred.

The anthraquinone process is a cyclic process, comprising ahydrogenation stage, where stream S3 is reacted with working solution inthe presence of a hydrogenation catalyst to convert at least part of thequinone to the corresponding hydroquinone, a subsequent oxidation stage,where the hydrogenated working solution containing hydroquinone isreacted with oxygen to form hydrogen peroxide and quinone, and anextraction stage, where hydrogen peroxide is extracted from the oxidizedworking solution with water to provide stream S6 as an aqueous solutionof hydrogen peroxide, with the extracted working solution being returnedto the hydrogenation stage to complete a reaction cycle.

In the hydrogenation stage, the working solution is reacted with streamS3 in the presence of a heterogeneous hydrogenation catalyst. During thereaction all or a part of the quinones are converted to thecorresponding hydroquinones. All hydrogenation catalysts known from theprior art for the anthraquinone cyclic process can be used as catalystsin the hydrogenation stage. Noble metal catalysts containing palladiumas the principal component are preferred. The catalysts can be used as afixed bed catalysts or as a suspended catalyst and suspended catalystscan be either unsupported catalysts, such as palladium black, orsupported catalysts, with suspended supported catalysts being preferred.SiO₂, TiO₂, Al₂O₃ and mixed oxides thereof, as well as zeolites, BaSO₄or polysiloxanes, are can be used as support materials for fixed-bedcatalysts or supported suspended catalysts, with TiO₂ and SiO₂/TiO₂mixed oxides being preferred. Catalysts in the form of monolithic orhoneycombed moldings, the surface of which is coated with the noblemetal, can also be used. Hydrogenation can be carried out instirred-tank reactors, tube reactors, fixed-bed reactors, loop reactorsor air-lift reactors which can be equipped with devices for distributingstream S3 in the working solution, such as static mixers or injectionnozzles. Preferably, a tube reactor with a recycle and a Venturi nozzlefor injecting stream S3 into the reactor feed as known from WO 02/34668is used. Hydrogenation is carried out at a temperature of from 20 to100° C., preferably 45 to 75° C., and a pressure of from 0.1 MPa to1MPa, preferably 0.2 MPa to 0.5 MPa. The hydrogenation is preferablyperformed in such a way that essentially all hydrogen introduced withstream S3 into the hydrogenation reactor is consumed in a single passthrough the reactor. The ratio between stream S3 and working solutionfed to the hydrogenation reactor is preferably chosen to convert between30 and 80% of the quinones to the corresponding hydroquinones. If amixture of 2-alkylanthraquinones and 2-alkyltetrahydroanthraquinones isused, the ratio between stream S3 and working solution is preferablychosen so that only the 2-alkyltetrahydroanthraquinones are converted tohydroquinones and the 2-alkylanthraquinones remain in the quinone form.

In the oxidation stage, the hydrogenated working solution from isreacted with an oxygen-containing gas, preferably with air or withoxygen enriched air. All oxidation reactors known from the prior art forthe anthraquinone process can be used for the oxidation, bubble columnsoperated in co-current being preferred. The bubble column can be freefrom internal devices, but preferably contains distribution devices inthe form of packings or sieve plates, most preferably sieve plates incombination with internal coolers. Oxidation is carried out at atemperature of from 30 to 70° C., preferably from 40 to 60° C. Oxidationis preferably performed with an excess of oxygen to convert more than90%, preferably more than 95%, of the hydroquinones to the quinone form.

In the extraction stage, the oxidized working solution containingdissolved hydrogen peroxide is extracted with an aqueous solution toprovide an aqueous hydrogen peroxide solution and an extracted oxidizedworking solution containing essentially no hydrogen peroxide. Deionizedwater, which may optionally contain additives for stabilizing hydrogenperoxide, adjusting the pH and/or corrosion protection, is preferablyused for extracting the hydrogen peroxide. Extraction is preferablycarried out in a counter-current continuous extraction column,sieve-plate columns being most preferred. The aqueous hydrogen peroxidesolution obtained by extraction may be used directly as stream S6 or maybe concentrated by distilling off water at reduced pressure to providestream S6. The aqueous hydrogen peroxide solution obtained by extractionmay also be purified, preferably by washing with a solvent, which ispreferably a solvent comprised in the working solution.

The anthraquinone process preferably comprises at least one additionalstage for regenerating the working solution, where by-products formed inthe process are converted back to quinones. Regeneration is carried outby treating hydrogenated working solution with alumina or sodiumhydroxide, preferably using a side stream to the cyclic process. Inaddition to regeneration of hydrogenated working solution, extractedoxidized working solution may be regenerated in a side stream usingalumina, sodium hydroxide or an organic amine. Suitable methods forregenerating the working solution on an anthraquinone process are knownfrom the prior art.

In step f) of the integrated process of the invention, a part or all ofstream S6 is reacted with propene in the presence of an epoxidationcatalyst, providing a stream S7 comprising propene oxide. Propene can beused as a technical mixture with propane, preferably containing from 0.1to 15 vol-% of propane. Both homogeneous and heterogeneous epoxidationcatalysts may be used in step f). Suitable epoxidation catalysts andreaction conditions for reacting stream S6 with propene to form propeneoxide are known from the prior art. Suitable homogeneous epoxidationcatalysts are manganese complexes with polydentate nitrogen ligands, inparticular 1,4,7-trimethyl-1,4,7-triazacyclononane ligands, as knownfrom WO 2011/063937. Other suitable homogeneous epoxidation catalystsare heteropolytungstates and heteropolymolybdates, in particularpolytungstophosphates, as known from U.S. Pat. No. 5,274,140. Suitableheterogeneous epoxidation catalysts are titanium zeolites containingtitanium atoms on silicon lattice positions.

Preferably, a titanium silicalite catalyst is used, preferably with anMFI or MEL crystal structure, and most preferably titanium silicalite-1with MFI structure as known from EP 0 100 119 A1, is used. The titaniumsilicalite catalyst is preferably employed as a shaped catalyst in theform of granules, extrudates or shaped bodies. For the forming processthe catalyst may contain 1 to 99% of a binder or carrier material, allbinders and carrier materials being suitable that do not react withhydrogen peroxide or with the epoxide under the reaction conditionsemployed for the epoxidation, silica being preferred as binder.Extrudates with a diameter of 1 to 5 mm are preferably used as fixed bedcatalysts. Epoxidation with a titanium silicalite catalyst is preferablycarried out in a methanol solvent. The methanol solvent can be atechnical grade methanol, a solvent stream recovered in the work-up ofthe epoxidation reaction mixture or a mixture of both. The epoxidationis preferably carried out at a temperature of 30 to 80° C., morepreferably at 40 to 60° C., and a pressure of from 0.5 to 5 MPa, morepreferably 1.5 to 3.5 MPa. The epoxidation is preferably carried out ina fixed bed reactor by passing a mixture comprising propene, hydrogenperoxide and methanol over the catalyst fixed bed. The fixed bed reactoris preferably equipped with cooling means and cooled with a liquidcooling medium. The temperature profile within this reactor ispreferably maintained such that the cooling medium temperature of thecooling means is at least 40° C. and the maximum temperature within thecatalyst bed is 60° C. at the most, preferably 55° C.

The epoxidation reaction mixture is preferably passed through thecatalyst bed in down flow mode, preferably with a superficial velocityfrom 1 to 100 m/h, more preferably 5 to 50 m/h, most preferred 5 to 30m/h. The superficial velocity is defined as the ratio of volume flowrate/cross section of the catalyst bed. Additionally it is preferred topass the reaction mixture through the catalyst bed with a liquid hourlyspace velocity (LHSV) from 1 to 20 h⁻1, preferably 1.3 to 15 h⁻¹. It isparticularly preferred to maintain the catalyst bed in a trickle bedstate during the epoxidation reaction. Suitable conditions formaintaining the trickle bed state during the epoxidation reaction aredisclosed in WO 02/085873 on page 8 line 23 to page 9 line 15. Propeneis preferably employed in excess relative to the hydrogen peroxide inorder to achieve high hydrogen peroxide conversion, the molar ratio ofpropene to hydrogen peroxide preferably being chosen in the range from1.1 to 30. The methanol solvent is preferably used in the epoxidation ina weight ratio of 0.5 to 20 relative to the amount of stream S6. Theamount of catalyst employed may be varied within wide limits and ispreferably chosen so that a hydrogen peroxide consumption of more than90%, preferably more than 95%, is achieved within 1 minute to 5 hoursunder the employed epoxidation reaction conditions. Most preferably, theepoxidation reaction is carried out with a catalyst fixed bed maintainedin a trickle bed state at a pressure close to the vapour pressure ofpropene at the reaction temperature, using an excess of propene thatprovides a reaction mixture comprising two liquid phases, a methanolrich phase and a propene rich liquid phase. Two or more fixed bedreactors may be operated in parallel or in series in order to be able tooperate the epoxidation process continuously when regenerating theepoxidation catalyst. Regeneration of the epoxidation catalyst can becarried out by calcination, by treatment with a heated gas, preferablyan oxygen containing gas or by a solvent wash, preferably by theperiodic regeneration described in WO 2005/000827.

In step g) of the integrated process of the invention, propene oxide isseparated from stream S7 provided in step f) as reaction product.Propene oxide may be separated from stream S7 by distillation orextractive distillation using methods known from the prior art.Preferably, propene oxide is separated from stream S7 by distillationafter a pressure release stage which removes most of the non-reactedpropene from stream S7. When a methanol solvent is used in step l andstream S7 comprises methanol, the distillation is preferably carried outin at least two columns, operating the first column to provide a crudepropene overhead product containing from 20 to 60% of the methanolcontained in stream S7 and further purifying the overhead product by atleast one additional distillation. The overhead product is preferablyfurther purified by distilling off remaining propene and propane,followed by extractive distillation, most preferably using theextractive distillation method of WO 2004/048355 for additional removalof carbonyl compounds.

If only a part of stream S4 obtained in step b) is reacted with analkanol in step c), the remainder of stream S4 can be used as a feed fora different reaction. Preferably, a feed stream S1 consistingessentially of iso-butane is used in step a) and the part of stream S4that is not fed to step b) is used as feed for a catalytic partialoxidation converting iso-butene to methacrylic acid. Oxidation ofiso-butene to methacrylic acid is preferably carried out in two gasphase oxidation stages via methacrolein as an intermediate. In the firststage, iso-butene is preferably reacted with oxygen at a temperature offrom 300 to 420° C. and a pressure of from 1 to 5 bar in the presence ofa molybdenum bismuth mixed oxide catalyst which can be promoted with atleast one of iron, cobalt, nickel, antimony, tungsten and alkali metals.Preferably, a catalyst having a composition as described in EP 267 556,page 3, lines 19 to 33 is used. The reaction is preferably carried outby passing a gas stream comprising iso-butene, oxygen, water and aninert gas with a volume ratio of iso-butene:O_(2:)H₂O:inert gas of1:0.5-5:1-20:3-30, preferably 1:1-3:2-10:7-20, over a catalyst fixedbed. The inert gas is preferably nitrogen, carbon dioxide, methane or amixture of nitrogen and carbon dioxide. In the second stage,methacrolein is preferably reacted with oxygen at a temperature of from250 to 350° C. and a pressure of from 1 to 3 bar in the presence of aphosphomolybdic acid catalyst promoted with an alkali metal, which canbe further promoted with at least one of copper, vanadium, arsenic andantimony. Preferably, a catalyst having a composition as described in EP376 117, page 2, line 50 to page 3, line 21 is used. The catalyst ispreferably employed as a fixed bed. The reaction is preferably carriedout by passing a gas stream comprising methacrolein, oxygen, water andan inert gas with a volume ratio of iso-butene:O_(2:)H₂O:inert gas of1:1-5:2 20:3-30, preferably 1:1-4:3-10:7-18, over a catalyst fixed bed.In a preferred embodiment, the reaction product stream from the firststage is fed directly as feed to the second stage, optionally withadding further oxygen and/or inert gas. Methacrylic acid andmethacrolein are separated from the reaction product gas of the secondstage, preferably by condensation, absorption or a combination of both,and the separated methacrolein is preferably returned to the entry ofthe second stage. Residual gas remaining after separation of methacrylicacid and methacrolein may be subjected to combustion and the combustiongas may be returned to the entry of the first or second stage to provideall or part of the inert gas.

In one embodiment of the integrated process of the invention, feedstream S1 to step a) comprises n-butane. The dehydrogenation of n-butanein step a) will then generate 1-butene, cis-2-butene, trans-2-butene and1,3-butadiene and stream S4 will comprise these compounds in addition toiso-butene. 1,3-butadiene is preferably removed from stream S4 beforereacting stream S4 in step c). 1,3-butadiene can be removed from streamS4 by selective extraction, preferably using N-methylpyrrolidone, mostpreferably as described in U.S. Pat. No. 6,337,429. Alternatively,1,3-butadiene can be removed from stream S4 by partial hydrogenation,preferably using the method of DE 31 43 647 and the reactorconfiguration of DE 195 24 971. In a further preferred embodiment, astream S8 comprising 1-butene is separated in step d) and from thisstream S8 a stream S10 consisting essentially of 1-butene is separated.Stream S8 preferably consists essentially of C4 hydrocarbons and can beobtained from stream S5 by separating the alkyl tert-butyl ether andnon-reacted alkanol by known methods, such as distillation, extractivedistillation or extraction, as described further above. Stream S8 can beseparated in a first rectification step into an overhead productconsisting essentially of 1-butene, isobutene and lower boilinghydrocarbons and a bottoms product, containing essentially all n-butane,cis-2-butene and trans-2-butene fed with stream S8. The firstrectification step can be carried out in a rectification column havingmore than 100 and preferably 150 to 200 theoretical plates, operated ata pressure of 0.4 to 1.0 MPa, preferably 0.5 to 0.7 MPa and atemperature of from 35 to 80° C., preferably 40 to 65° C. The overheadproduct from the first rectification step can be separated in a secondrectification step into a an overhead product containing isobutene andlower boiling hydrocarbons, which can be recycled as stream S9 to stepa), and a bottoms product consisting essentially of 1-butene as streamS10. The second rectification step can be carried out in a rectificationcolumn having more than 100 and preferably 150 to 200 theoreticalplates, operated at a reflux ratio of from 30 to 60. Stream S10 can beused as a comonomer in the preparation of polyethylene (LLDPE or HDPE)and ethylene-propylene copolymers or starting material for thepreparation of 2-butanol, butene oxide or valeraldehyde. The bottomsproduct from the first rectification step can be used as feed for anolefin oligomerization to prepare olefins having eight, twelve, sixteen,twenty or more carbon atoms, which have use as starting materials formaking C9 and C13 alcohols for the manufacture of plasticisers, C13, C17and C21 alcohols for the manufacture of detergents and high purity C16+paraffins. A suitable methods for olefin oligomerization is known fromHydrocarbon Process., Int. Ed. (1986) 65 (2. Sect.1), pages 31-33 as theOCTOL process.

FIG. 1 shows a flow chart for an embodiment of the integrated process ofthe invention where a feed stream S1 (1) comprising n-butane in additionto iso-butane is fed to the dehydrogenating step (2). Dehydrogenationprovides a stream S2 (3) comprising iso-butene and hydrogen. Due to thepresence of n-butane in stream S1, stream S2 also comprises the linearbutenes 1-butene, cis-2-butene and trans-2-butene. Stream S2 alsocomprises unreacted iso-butane and n-butane. Stream S2 (3) is separatedin a separation step (4) into a stream S3 (6) consisting essentially ofhydrogen and a stream S4 (5) comprising iso-butene. In addition toiso-butene, stream S4 also comprises the linear butenes and unreactediso-butane and n-butane. Stream S3 (6) consisting essentially ofhydrogen is reacted with oxygen (7) in step (8) to provide a stream S6(9) comprising hydrogen peroxide. Step (8) is preferably ananthraquinone process using air as oxygen source and providing stream S6as an aqueous hydrogen peroxide solution. Stream S6 (9) is reacted withpropene (10) in step (11) to provide a stream S7 (12) comprising propeneoxide. Step (11) is preferably carried out with a shaped titaniumsilicalite catalyst in a methanol solvent using a fixed bed reactor. Ina subsequent separation step (13), propene oxide (14) is separated fromStream S7 (12). Stream S4 (5) is reacted with an alkanol (15) in step(16) in the presence of a solid acid catalyst, providing a stream S5(17) comprising an alkyl tert-butyl ether. The alkanol is preferablymethanol, providing a stream S5 (17) comprising methyl tert-butyl ether.In a subsequent separation step (18), the alkyl tert-butyl ether (20) isseparated from stream S5 (17) and a stream S8 (19) comprising unreactediso-butane and alkanol is obtained. In addition to unreacted iso-butane,stream S8 also contains unreacted n-butane and the linear butenes. In afurther separation step (21), stream S8 (19) is further separated into astream S9 (22) comprising more than 80% by weight iso-butane, which isrecycled to the dehydrogenating step (2), a stream stream S10 (23)consisting essentially of 1-butene, and one or more streams, not shownin FIG. 1, which comprise alkanol, n-butane, cis-2-butene andtrans-2-butene.

FIG. 2 shows a flow chart for an embodiment of the integrated process ofthe invention additionally comprising a step of isomerizing n-butane toiso-butane to provide the feed stream to step a). A feed stream (24)comprising n-butane is fed to a step (25) of isomerizing n-butane toiso-butane to provide a mixture (26) comprising n-butane and iso-butane.Preferably, feed stream (24) essentially consists of n-butane andmixture (26) comprises n-butane and iso-butane at the equilibrium ratioestablished at the temperature used in step (25). In a subsequentseparation step (27), n-butane is separated from mixture (26) andrecycled to step (25) of isomerizing n-butane to iso-butane, providingfeed stream S1 (1) comprising iso-butane. In this embodiment, n-butaneis preferably separated as complete as possible in step (27) in order toprovide streams S4 (5), S5 (17) and S8 (19) that contain essentially non-butane and no linear butenes.

The integrated process of the invention saves raw material compared toseparate processes for making alkyl tert-butyl ether and propene oxideby using hydrogen obtained by dehydrogenating iso-butane for makinghydrogen peroxide for use in the preparation of propene oxide. Comparedto the Halcon process of coupled production of MTBE and propene oxide,the integrated process of the invention has the advantage of providingmuch better flexibility in varying the production ratio of tert-butylether and propene oxide. If a feed stream comprising n-butane inaddition to iso-butane is used, the integrated process of the inventionalso provides 1-butene as a valuable product. In the embodimentcomprising a step of isomerizing n-butane to iso-butane, the productionratios of tert-butyl ether, 1-butene and propene oxide can be variedeven more flexibly by varying the ratio of iso-butane to n-butane infeed stream S1.

List of Reference Signs

-   1 Feed stream S1 comprising iso-butane-   2 Step of dehydrogenating iso-butane-   3 Stream S2 comprising iso-butene and hydrogen-   4 Step of separating iso-butene and hydrogen-   5 Stream S4 comprising iso-butene-   6 Stream S3 consisting essentially of hydrogen-   7 Oxygen-   8 Step of reacting hydrogen with oxygen to provide hydrogen peroxide-   9 Stream S6 comprising hydrogen peroxide-   10 Propene-   11 Step of reacting hydrogen peroxide with propene to provide    propene oxide-   12 Stream S7 comprising propene oxide-   13 Step of separating propene oxide-   14 Propene oxide-   15 Alkanol-   16 Step of reacting iso-butene with alkanol-   17 Stream S5 comprising alkyl tert-butyl ether-   18 Step of separating alkyl tert-butyl ether-   19 Stream S8 comprising unreacted iso-butane-   20 Alkyl tert-butyl ether-   21 Step of separating unreacted iso-butane-   22 Stream S9 comprising more than 80% by weight iso-butane-   23 Stream S10 consisting essentially of 1-butene-   24 Feed stream comprising n-butane-   25 Step of isomerizing n-butane to iso-butane-   26 Stream comprising n-butane and iso-butane-   27 Step of separating n-butane and iso-butane-   28 Recycle stream comprising n-butane

1-9. (canceled)
 10. An integrated process for making propene oxide andan alkyl tert-butyl ether, comprising: a) a step of dehydrogenating afeed stream S1 comprising iso-butane, to provide a stream, S2,comprising iso-butene and hydrogen; b) a separation step separatingstream S2 into a stream, S3, consisting essentially of hydrogen and astream, S4, comprising iso-butene; c) a step of reacting a part or allof stream S4 with an alkanol in the presence of a solid acid catalyst,to provide a stream, S5, comprising an alkyl tert-butyl ether; d) aseparation step separating the alkyl tert-butyl ether from stream S5; e)a step of reacting a part or all of stream S3 with oxygen, to provide astream, S6, comprising hydrogen peroxide; f) a step of reacting a partor all of stream S6 with propene in the presence of an epoxidationcatalyst, to provide a stream, S7, comprising propene oxide; and g) aseparation step separating propene oxide from stream S7.
 11. The processof claim 10, wherein the alkanol is methanol.
 12. The process of claim10, wherein a stream, S8, comprising unreacted iso-butane is separatedin step d), a stream S9 comprising more than 80% by weight iso-butane isseparated from stream S8 and stream S9 is recycled to step a).
 13. Theprocess of claim 10, wherein the feed stream S1 of step a) comprisesn-butane, a stream, S8, comprising 1-butene is separated in step d) anda stream, S10, consisting essentially of 1-butene is separated from saidstream S8.
 14. The process of claim 10, wherein the feed stream S1 tostep a) comprises n-butane and wherein 1,3-butadiene is removed fromstream S4 before reacting it in step c).
 15. The process of claim 10,comprising an additional step of isomerizing n-butane to iso-butane, toprovide the feed stream S1 to step a).
 16. The process of claim 10,wherein step e) is carried out as an anthraquinone process providing a20 to 75% by weight aqueous solution of hydrogen peroxide as stream S6.17. The process of claim 10, wherein a titanium silicalite catalyst isused as epoxidation catalyst in step f).
 18. The process of claim 10,wherein additional hydrogen is fed to step e).
 19. An integrated processfor making propene oxide and methyl tert-butyl ether, comprising: a) astep of dehydrogenating a feed stream, S1, comprising iso-butane, toprovide a stream, S2, comprising iso-butene and hydrogen; b) aseparation step separating stream S2 into a stream, S3, consistingessentially of hydrogen and a stream, S4, comprising iso-butene; c) astep of reacting a part or all of stream S4 with methanol in thepresence of a solid acid catalyst, to provide a stream, S5, comprisingmethyl tert-butyl ether; d) a separation step separating methyltert-butyl ether from stream S5; e) a step of reacting a part or all ofstream S3 with oxygen in an anthraquinone process, to provide a stream,S6, comprising hydrogen peroxide as a 20 to 75% by weight aqueoussolution of hydrogen peroxide; f) a step of reacting a part or all ofstream S6 with propene in the presence of an epoxidation catalyst, toprovide a stream, S7, comprising propene oxide; and g) a separation stepseparating propene oxide from stream S7.
 20. The process of claim 19,wherein a titanium silicalite catalyst is used as epoxidation catalystin step f).
 21. The process of claim 19, wherein the feed stream S1 tostep a) comprises n-butane, a stream, S8, comprising 1-butene isseparated in step d) and a stream, S10, consisting essentially of1-butene is separated from said stream S8.
 22. The process of claim 21,wherein 1,3-butadiene is removed from stream S4 before reacting it instep c).